Isoforming



3 Sheets-Sheet l Aug. 1o, 1943. E. W. @HELE ET AL ISOFORMING Filed Nov'. 28, 1940 Aug 10, 1943- E. w. THIELE ET A1. 2,326,705

ISOFORMING Filed Nov. 28, 1940 3 Sheets-Sheet 2 /3Z` :"f-ff* REGENEP/lr/NG N cmq/HEER,

Allg 10, 1943 u E. w. THIELE ET AL 2,326,705

ISOFORMING Filed Nov. 28, 1940 3 Sheets-Sheet 5 511A CE mzacff'yfV/V/HR) 12 3 2 Patented Aug. 10, 1943 ISOFORMING Ernest W. Th-iele, George E. Schmitkons, and

Carl Max Hull, Chicago, Ill., assignors to Standard Oil Company, Chicago, Ill., al corporation of Indiana Application November 28, 1940, Serial No. 367,582

11 Claims.

This invention relates to an improved method of making high antiknock motor fuels from charging stocks consisting essentially of gas oils' and heavier hydrocarbons.

Heretofore, most motor fuel has been produced by thermal cracking and most reneries are, therefore, equipped with thermal cracking systems'. Demands for higher octane number gasoline have given rise to a new problem because the maximum octane number obtainable by cornmercial thermal' cracking is about 65 to 74 CFR-M. Furthermore, thermally cracked naphtha is` not suiciently responsive to tetraethyl lead to make it commercially feasible to obtain the desired higher octane numbers by this route. Hence reners have turned to expensive and complicated processes of catalytic cracking, destructive hydrogenation, hydroforming, aromatization, alkylation, isomerization, polymerization, etc. We have discovered that the problem can be solved most advantageously and economically by subjecting thermally cracked naphtha to a simple high temperature contacting with a catalyst to eiect what we term isoforming Isoformlng is distinctly diierent from all prior art processes in that it employs thermally cracked naphtha as its charging stock and in that it produces 95% to 99% yields based on this charging stock with surprisingly low losses to gas, coke and heavier-than-gasoline hydrocarbons.

In catalytic reforming and cracking processes the yield-octane curve tends to flatten out, but

but always with yields considerably below 100% and with considerable losses to gas, coke and heavier than gasoline hydrocarbons. Hydrogenation, aromatization and hydroforming have been proposed for increasing the octane number of naphtha but experience has shown that thermally cracked naphtha is not particularly responsive to such processes, that only a very small improvement is obtained in octane number and that losses are much higher than those obtainable by isoforming. Innumerable complicated and expensive processes have been proposed in an effort to solve this problem and'since it appears to be practically unsolvable, many refiners are electing to change their refining processes, to substitute catalytic cracking, destructive hydrogenat-ion, etc. in order to meet the demands for higher octane number gasoline. Our isoforming process makes it possible to utilize existing thermal cracking equipment and to meet octane requirements with minimum `losses to gas and carbon and with a catalyst holding time, i. e., time between regenerations, which .far exceeds the catalyst holding time possible in catalytic cracking processes. An important feature of the In addition to obtaining incre-ases of 5 to 15 it does not bend backwards; in isoforming we have found that yield-octane curve actually does bend backwards, showing a definite optimum octane number with a yield of 94 to 99%, the maximum in most cases being with a yield of about 96% or 97%. Also in catalytic reforming and cracking the octane number is gradually lowered with increasing space velocities (shorter times of Contact), while in the isoforming reaction there is a definite peak inthe curve. At atmospheric pressures and with a fair catalyst this peak usually is within the range of 4 to 40, usually about 12 volumes of liquid thermally cracked naphtha per volume of catalyst space per hour.

Thermally cracked naphtha has been contacted with clay for improving its stability against gum formation, for lowering its sulphur content, etc. but the conditions of these clay contacting processes have been such that very little if any improvement in octane numbers was accomplished. Both thermal and catalytic reforming or isomerization have been applied to virgin naphtha,

octane numbers at a relatively high octane number level, we obtain the advantage of increased responsiveness to lead tetraethyl.

Isoformng has practically no effect on catalytcally cracked naphtha. We have discovered, however, that the gas oil `from a catalytic cracking process may be thermally cracked to give a thermally cracked naphtha which does respond to the isoformlng treatment.. Thermal cracking evidently ruptures the molecules to form a different type of naphtha than is formed by catalytic processes. The initial thermal cracking step is thus an essential element of our combined process. The thermal cracking should preferably be eiected at relatively low pressures and high temperatures although pressures of 300 to 750 pounds or even higher pressures may be advantageously used. The thermal cracking may be carried out on a once-through or on a recycle basis. Our invention is applicable lto all types of thermally cracked naphtha. The expression thermally cracked naphtha as used in this specification refers only to naphtha in the gasoline boiling range produced by the thermal cracking of gas oil and other hydrocarbon charging stocks which boil above the gasoline boiling range. The species lof the invention hereinafter described will be was pointed out that isoforming could be eilected in a powdered catalyst system. In the prior application the conditions were pointed out for isoforming in a fixed bed system wherein the reaction waseffected at a pressure oi' about atmospheric to 50 pounds, one exemple being 2 atmospheres. It was pointed out that the space velocity was of a higher order than employed in catalytic cracking or aromatization, being about 8 to 40 v/v/hr., a particular example being 12 v/v/hr. It was pointed out that run lengths between regenerations might be between 8 to 24 hours. Our present invention relates to an improved method and means for effecting this same reaction in a powdered catalyst system.

In our powdered catalyst system the catalyst is maintained in aerated form and may thus be handled as a fluid. Due to the abrasive character of the powdered catalyst it is desirable that it be conveyed through the system by pneumatic means rather than by the use of mechanical pumps. This can be readily accomplished by `elevating the catalyst by the gas lift effect of gases and vapors in the system. The necessary pressures for introducing catalyst into fluid streams may be obtained by maintaining a dense aerated catalyst in a standpipe. It has been found that with slight aeration'the powdered catalyst settles into a dense phase which behaves like a liquid.

In the powdered catalyst system the cracked naphtha is heated to about the same temperature as in the fixed bed system. Hot powdered catalyst is then introduced into the hot vapor stream and passed through a reaction chamber or tube. The horizontal cross-sectional area of this chamber or tube is so proportioned with respect to the amount and velocity of incoming vapors and the relative amount of catalyst in said vapors that the necessary time of contact is obtained with the required amount of catalyst. This contact may be obtained either under dense phase or light phase conditions, but in either case the vapors should contact about the same amount of catalyst as it would contact in a iixed bed system.

Under dense phase conditions a reaction chamber of relatively large diameter is employed and although the vapors leaving the top of the reactor withdraw with them the same amount of catalyst as is introduced with the incoming va,- pors, the actual amount of catalyst so introduced and withdrawn is very small compared with the amount of vapors introduced and withdrawn. AA

be dependent upon the catalyst-to-oil ratio employed, the carbon tolerance of the catalyst, the reactor design, etc.

Under light phase conditions the reactor may be of small cross-sectional area and a relatively large amount of catalyst is introduced into the hot cracked naphtha vapors and carried thereby through the reactor without appreciable settling. Here the amount of catalyst in thereactor will be substantially the same as in the dense phase operation, but the holding time in the reactor may be a matter of seconds instead of hours. y

In either the dense phase or light phase operations, catalyst is separated from reaction vapors by centrifugal separators and the vapors are then fractionated. The small amount of catalyst which remains in the vapors may be recovered by means of the Cottrell precipitators or may be recovered with condensed vapor and recycled either to the isoforming step or to the thermal cracking step.

The catalyst is stripped of hydrocarbons and passed through a regeneration chamber wherein carbonaceous material is burned from the spent catalyst, preferably while said catalyst is in a dense aerated state. Here again the aerated catalyst behaves as a liquid and the introduced oxidizing gas serves to agitate this dense phase and give it the appearance of boiling, an amount of regenerated catalyst being withdrawn from the top of the chamber equivalent to that which is being constantly introduced thereto. An mportant feature of this regeneration system is the fact that regeneration temperature is substantially constant throughout the entire chamber, i. e., there are no hot spots. Another feature is that regeneration temperature may be controlled by continuously withdrawing some of the aerated catalyst through a heat exchanger and reintroducing it into the regeneration chamber or by simply passing a cooling iiuid through coils in the zone of the aerated catalyst undergoing regeneration. Alternatively the temperature may be controlled by separating catalyst from the gases leaving the top of the regeneration chamber, cooling vthe separated catalyst andreturning it to the regeneration chamber. After removal of regeneration gases from the regenerated catalyst in cyclone separators and the like, this catalyst is reintroduced into the hot cracked naphtha stream for further isoforming.

The invention will be more clearly understood from the following detailed description and from the accompanying drawings which form a part of this speciilcation and in which Figure 1 is a diagrammatic iiow sheet of our process employing dense phase isoforming with powdered catalyst in a fluid system.

Figure 2 is a diagrammatic flow sheet o f light phase powdered catalyst isoforming.

Figure 3 is a chart showing the unusual yieldoctane relationships that are characteristic of the isoformingprocess, and

Figure 4 is a chart showing the unusual relationship between octane number and space velocity in the catalyst chamber.

The charging stock to our system may be a Mid-Continent gas oil although it may comprise a gas oil from any other source and it may include any hydrocarbons heavier than gasoline, i. e., reduced crudes, residual stock, etc. Any conventional thermal cracking process may be employed for preparing the feed stock for isoforming.

i In Figure 1, we have dlagrammaticallyillustrated thermal cracking of the continuous pressure still type wherein Mid-Continent gas oil is forced by `pump through line to coils I2 of pipe still furnace I3 under a pressure of about 300 pounds per square inch and a temperature 'of about 925 F. The thermally cracked products are then introduced by transfer line I4 either directly or through a conventional soaking drum (not shown) to evaporator I5 which is provided with a tar draw-od line I6 at the bottom. Cracked naphtha and gases are taken overhead through line I1 leading to a bubble tower I8. A pressure reducing valve I 9 may be employed in line I 1 if it is desired to obtain the fractionation in the bubble tower at lower than reaction pressures. A reboiler 20 at the base of the bubble tower insures the removal of all of the naphtha from the gas oil which is withdrawn from the base of the tower through line 2| and which may be withdrawn from the system through line 22 or recycled by pump 23 and line 24 to line II for further cracking.

The naphthas and lighter hydrocarbons are taken overhead through line 25 and cooler 26 and are then introduced into a gas separator or receiver tank 2l from which gases are vented through line 28 and liquid naphtha is withdrawn through line 29. A portion of the naphtha is recycled through line 30 by pump 3|,for reflux in top of bubble tower I8. The remainder of the napththa or a fraction thereof is introduced through line 32 to theisoforming step of our process. A feature of our invention is the fact that the thermally cracked naphtha does not have to be stabilized before it is introduced to the isoforming system. In fact the total overhead products from tower I3 may be introduced directly t0 line 32 through line 25 when no part of it is required for reflux in tower I8. Alternatively, cooler 26 may condense only the liquid required for reux and all of the naphtha and vapors may pass through lines 28 and 25 to line 32 for charging to the isoforming process.

The thermally cracked naphtha for our isoforming process should have an end point of about 400 to 450 F. It should be substantially free from gas oil since such material tends to foul the catalyst and to interfere with the proper operation of the isoforming process. The boiling range of the thermally cracked naphtha may be relatively wide but the greatest improvement is eiected on the high boiling portions.

100 F. and in fact normallygaseous hydrocarbons may be beneficial in reducing the partial pressure of the thermally cracked naphtha under-z goingthe isoforming treatment. when normally gaseous hydrocarbons are -employed with the naphtha the overall pressures may be superatmospheric while the effective pressure is only atmospheric over even sub-atmospheric.

The thermally cracked naphtha produced as hereinabove described may have an olefin content of about 25-70%, an octane number of about 60 to '70 CFR- M and an A. P. I. gravity of about 50-60. A feature of our process is the octane improvement obtained at this relatively high octane level.

In one embodiment of our process this thermally cracked naphtha feed stock is introduced by pump 33 through coils 34 of pipe still furnace 35 and discharged therefrom through transfer line 36. The transfer line temperature should be The lower limit of this boiling range may be lower than` about 350y to 1025 F., preferably about 925 to y 975 F. and its pressure should be about from atmospheric to 50 pounds per square inch, preferably about 10 to 15 pounds per square inch. It should be understood, however, that higher pressures may be employed and that lower eiective pressures may be obtained by the 'use of hydrocarbon gases, steam, etc.

Catalyst may be injected into this hot thermally cracked naphtha stream through line 99 to obtain in transfer line a catalyst-to-oil weight ratio of about .001:1 to .1:1, preferably about .006:1 to .02:1 for example .01:1. -Alternatively, catalyst may be introduced through line 90 to the stock entering the pipestill, particularly when the amount of catalyst so introduced is relatively small. Steam is ,preferably introduced through line 01 or line 98 for injecting the catalyst into the oil stream and the amount of steam so used may be about 1 to 15% by weight, for example about 10%, based on oil charge. 'I'he powdered catalyst is carried by the steam and hot cracked naphtha vapors through line 36 to the base of isoforming reaction chamber |00.

The reaction chamber for a 20,000 barrels per day plant may be about 18 to 20 feet in diameter and about 15 to 20 feet high to allow a dense phase of about 5 or more feet in depth and of about 1200 to 3000 cubic feet in volume. The incoming stream is so introduced as to maintain this dense phase of aerated catalyst in the lower part of the reaction chamber and a light phase in the upper part of said chamber. The upward vapor velocity for this purpose may be maintained at about 1 or 11/2 feet per second with catalyst having an average particle size of 20 to 80 microns but may be higher with coarser catalyst material and lower with a i'lner catalyst powder. Suitable distributors may be used at the base of the reactor to insure uniform distribution.

The depth of the densephase ln the reactor may be determined by the use of a simple man- In other Words,

ometer indicating the difference in pressure between upper and lower pointsl in the catalyst chamber. The depth of the dense layer may be increased by decreasing the linear upward velocity of vapors in the reactor or by increasing the catalyst-to-oil ratio and either or both expedients may be employed for maintaining the desired catalyst residence time in the reactor.

The isoforming reaction products with suspended catalyst are withdrawn from the top of the reaction chamber through line l0! and introduced into cyclone separator |02 for removing the bulk of the catalyst therefrom, the catalyst being passed by line |03 into an enlarged section of a stripping tower |04 which may be at a pressure of about 6 pounds per square inch. Vapors from cyclone separator |02 are then passed through line |05 to a second cyclone separator |06 from which additional catalyst material is returned to tower |00 through line |01.- `Vapors from separator |05 may be passed through a Cottrell precipitator or through further cyclone separators, or through any other conventionalmeans (not shown) for separating the last traces of catalyst from the vapors. These vapors are then passed through line 00 to the fractionation system.

The spent catalyst which is introduced into stripping tower |04 passes over bales |00, thence to a standpipe section |09 which is of much smaller cross-sectional area than stripping section |00. The stand-pipe section is preferably about 20 to 30 feet high and it provides a static head (partly due to the pressure in stripper |04) of about 10 to 15 pounds per square inch. An inert gas such as steam or light hydrocarbon gases is introduced at the base of the standpipe through line H at a rate sufficient to maintain the catalyst in dense phase aerated condition. This inert gas serves as a stripping iluid in the enlarged stripping section |04 although additional stripping fluid may be introduced at the base of this enlarged section and at other points if desired.

Aerated catalyst is passed in amounts regulated by slide valve I|| into line ||2 from which it is forced by air from line ||3 to catalyst regeneration chamber ||4 which may operate at about 5 pounds perV square inch. Since the catalyst may be at a temperature of 750 to 900 F. at this point a certain amount of combustion will be effected in line I I2 but the amount of air introduced at this point is insuicient to cause the temperature to exceed about 1100" F. Additional air is introduced at the base of the regeneration chamber through line ||5 in amounts suillcient to burn the necessary amount of carbonaceous material from the catalyst. The cross-sectional area of the regeneration chamber is so proportioned that upward velocity of the regeneration gases maintains a dense catalyst phase the depth of which is preferably from about one-half tothree-fourths' the height of the regeneration chamber.

A feature of this regeneration system is that a substantially constant temperature prevails throughout the dense phase in the regeneration chamber. Since the dense Iphase is behaving like a boiling liquid its temperature may be controlled by conventional means employed in liquid systems, i. e., by passing a cooling fluid through coil I I6 immersed in this dense phase or by recycling the aerated dense phase catalyst material through an external heat exchanger. Alternatively, the temperature may be regulated by admixing relatively cold catalyst of low carbon content with the catalyst which is to undergo regeneration so' that the heat generated by the combustion of carbonaceous deposits increases the sensible heat in this additional catalyst.

Regeneration gases leaving the regeneration chamber through line lI'I carry away regenerated catalyst as fast as spent catalyst is introduced to line 'I|2. The catalyst-laden regeneration gases are passed to cyclone separator I I8 from as steam or iiue gas through line |21. This gas actsl as a stripping gas in enlarged section |20 although additional stripping gas may be introduced at the base of this section and at other points ii' needed. The stripping gas is withdrawn through line |28" and any catalyst carried thereby may be removed by means of Cottrell precipitators, etc. as hereinabove described.

Aerated catalyst from the base of standpipe |26 is introduced in amounts regulated by slide valve |29 into injector vessel |30 from which it is injected by steam introduced through line 91 into lines 99 and 36 as hereinabove described.

Vapors from the isoforming catalyst chamber together with vapors recovered from spent catalyst in stripping chamber |04 and discharged therefrom through line |3| are passed through line to fractionator tower 4| which is provided with suitable reboiler means 42 at the base thereof. The amount of heavier-than-gasoline hydrocarbons removed from the base of fractionator 4| through line 43 is extremely small if cracked naphtha of 400 F. end point or lower is fed to reactor |00 and such heavier-than-gasoline hydrocarbons may, therefore, be removed either continuously or intermittently and withdrawn from the system through line 44 or recycled by pump 45 through line 46 to line Il for thermal cracking or to line 32 for further treatment in the isoforining process.

End point isoformate together with hydrocarbon gases is taken overhead from tower 4I through line 41 and cooler 48 to receiver 49. Uncondensed gases are withdrawn from receiver 49 through line 50 and they may be compressed or absorbed in oil to recover gasoline components by `procedures familiar to the petroleum industry.

which the bulk of the catalyst is returned through i line H9 to enlarged stripper section |20 which may -be at a pressure of about 1 pound per square inch. The gases from H8 are passed through line I 2| to a second cyclone separator |22 from which further amounts of catalyst are returned to stripper |20 through line |23. Any number of cyclone separators may thus be employed and if desired auxiliary separation means such as a Pease-Anthony scrubber or a Cottrell precipitator (not shown) may be employed for removing the iinal traces of catalyst from regeneration gases which are Withdrawn through line |24. The heat contained in these regeneration gases may be used for generating steam or for any other purpose.

After passing over bailies |25 in stripper section |20, the regenerated catalyst passes into standpipe |26 which may be about 60 to 70 feet in height in order to obtain a fluid head at the base thereof of about 15 lpounds per square inch. The

A part of the liquid from receiver 49 is introduced by pump 5| and line 52 into the top of tower 4I to serve as reux. The rest of this liquid is introduced into stabilizer system 53 from which propane and lighter gases are withdrawn through line 54 'and finished isoformate is Withdrawn through line 55.

In the example hereinabove described, with a transfer line temperature of about 950 to 975 F., the isoforming is effected at a temperature of about 925 to 950 F., the average temperature in the isoforming reaction chamber being about 25 F. lower than the transfer line temperature. The catalyst-to-oil weight ratio in the stream entering the reaction chamber is about .006:1 to .02:1 for example about .01:1. By maintaining a dense phase in this reaction chamber of about 5 to 15 feet in depth 'and an upward vapor Velocity of about 1 foot per second, the vapors are contacted with about the same amount of catalyst as would be the case in a fixed bed system of about 400 cubic feet. The density of a fixed bed is about 40 to 45 pounds per cubic foot. The density of powdered catalysts in standpipes may be about 25 pounds per cubic foot. In the reactor with a vapor velocity of about 1 foot per second and with catalyst particle size of about 20 to 80 microns the dense phase may be about 10 or 15 Apounds per cubic foot so that the dense phase volume should be at least 2 to 3 times the volume of a iixed bed in order that the vapors may contact an equivalent amount of catalyst in their passage through the reactor. The holding time of the powdered catalyst in the reactor may vary throughout a wide range, but unless catalyst is to be recycled without regeneration we prefer a dense phase catalyst holding time in the reactor of about 4 to 36 hours, preferably about 8 to 18 hours. For a holding time of 4 hours the catalyst-to-oil ratio may be about .02 and for a holding time of 12 hours it may be about .006 in the example hereinabove described (for a 20,000 barrel per day plant) An isoformate yield of about 97 to 99% is thus obtained with a coke yield of less than .1% and a dry gas yield of only slightly more than 1% and a heavier-than-gasoline material yield of about 1% or less. The improvement in knock rating of the resulting isoformate'over the knock rating of the cracked naphtha charging stock will vary somewhat with the charging stock but will be of the order of about 4 to 15 octane numbers, usually about 5 to 10 octane numbers. Coke still naphtha will show a greater octane improvement than naphtha resulting from thermal cracking which has been eiected under high pressure. The resulting isoformate shows an increased response to tetraethyl lead and its actual performance in automobile or aviation engines is farA superior to the performance obtainable by simply adding the tetraethyl fluid to the original cracked naphtha charging stock.

Another embodiment of our invention is illustrated in Figure 2 wherein the isoforming is effected by means of a relatively large catalyst-tooil ratio and wherein there is little or no phase separationin the reactor. The oil from transfer line 3E is introduced into reactor tube |32`in admixture with powdered catalyst from line |3i which is injected into the reactor tube my means of steam from line 97. 1f it is desired to avoid the use of steam, hot hydrocarbon gases or vapors may be passed through line 35a and serve as the injector for introducing powdered catalyst into reactor |32.

In this case catalyst-to-oil ratios may be from 1:1 to 10:1 or more and the vapor velocity in the reactor is suiicient to carry the powdered catalyst therethrough without appreciable settling. The size and shape of the reactor or reactors in this case should be so designed that the linear velocity therethrough will be about 5 to 30 feet per second and their total volume should be suicient to contain the same amount of catalyst as was used in the dense phase operation. Although the amount of catalyst in the reactor at any instant is roughly equivalent to the amount of catalyst in the dense phase of the previous example, the catalyst holding time on each passage through the reactor may be a matter of seconds instead of hours. In this system, therefore, the catalyst is not spent when it reaches the end of the reaction zone.

As in the previous example, the catalyst is separated from reaction vapors in cyclone separator |33, the catalyst being returned by line |38 to stripper column |35. Vapors from separator |33 are passed by line |36 to cyclone separator |31 from which catalyst is returned to the stripper column through line |38. Vapors are removed through line 40 for fractionation as hereinabove described.

After passing over baiiles |39 the catalyst is introduced into the standpipe from which it is returned through leg |4l| in amounts regulated by slide valve |42 to line |3| for reintroduction into the vapor stream.

Through another leg |53 of the standpipe a" small amount of catalyst is withdrawn in amounts regulated by valve |55 through line |45 wherein it is forced by air introduced by line |46 into regeneration zone |457. Additional air` for the regeneration may be introduced by line |68. The regenerated catalyst is separated from vapors in cyclone separators |49 and |50 and returned through stripper column |5| and line |52 to the upper part of standpipe |40. Inert stripping g'as such as flue gas or steam is introduced in line |53 for insuring the removal of oxygen from the catalyst, the stripping gas and the regeneration gas being vented through line |50. As in the previous example, the catalyst in standpipe MD is maintained in mobile liquid like form by introducing an inert aerating gas through line |55 immediately above slide valves |42 and |44. In this example as in the previous example the vapors are treated with approximately the same I amount of catalyst as with the fixed bed system the catalyst is regenerated before it is returnedl to the reactor while in the light or vapor phase system most of the catalyst is recycled and onl,v a small amount is regenerated during each cycle.

` I t should be understood that various intermediate systems may be used herein but in all cases the vapors should be contacted with about the same amount .of catalyst in about the same amount oi time in the reaction chamber or tube.

Isoforming'is not primarily a cracking reaction as evidenced by the 97 to 99% volume percent of liquid gasoline yield. The remarkably high yield for the octane number improvement obtained and the increased susceptibility to tetraethyl lead place this process in a class by itself as far as isomerization is concerned since such results can only be obtained by the isomerization of a thermally cracked naphtha as hereinabove described and under operating conditions and with catalysts which will now be described in further detail.

The catalysts employed for isoforming are preferably the type generally employed for cracking virgin gas oils and heavier hydrocarbons to obtain high octane numbers. Activated hydrosilicate of alumina has been found to give excellent results. Such catalysts may .be prepared from acid treating bentonite. The catalyst may also be prepared by depositing alumina or other metaloxides on silica gel by impregnation with appropriate salts of the metals. Examples of such other metal oxides are copper, magnesium, beryllium, and thorium. Cadmium, titanium, manganese, zirconium, vanadium and cerium have shown slight activity in particular catalysts tested. A ball-milled 50-50 mixture of magnesium oxide and silica gave less coke and less gas than the preferred alumina-silica catalyst, but iii-likewise gave less octane number improvement. Catalysts .of the natural or synthetic zeolite type may be employed, preferably after sodium is displaced or leached out of the catalyst. Catalyst may be obtained by the treatment of blast furnace slag with hydrochloric acid followed by coagulation of the acid solution, washing and drying. Applicants are not herein claiming any ncvelty'in the catalyst per se but they do not employ catalysts of the dehydrogenation, hydroforming or aromatization type. Generally speaking, catalysts of the dehydrogenation or aromatizing type, such as bauxite, activated alumina, or oxides of ,certain metals (such as those of chromium, tungsten, nickel, or molybdenum) on alumina, are inferior and may be detrimental. Likewise, ordinary untreated clays are inferior in the isoforming process.

While ordinary cracking catalysts of the type that produce high octane numbers are also good isoforming catalysts it does not follow that all cracking catalysts are suitable for isoforming or vice versa. Silica 'gel is a cracking catalyst, but is not effective for isoforming. Boron phosphate is a good isoforming catalyst but it is not effective for catalytic cracking.

In powdered isoforming systems of the type hereinabove discussed the particle size of the catalyst is of importance in connection with vapor velocities and the cross-sectional areas of reactor, regenerator and standpipes. We prefer to employ catalyst having a uniform particle size. Catalyst particles smaller than about 10 microns are more diiilcult to recover in the cyclone separators than catalyst of about 50 microns. Catalyst particles exceeding about 80 or 90 microns require increased vapor velocities in order to maintain them in the dense phase aerated liquidlike state but are more easily separated from gases. It should be understood, of course, that with smaller average particle size lower space velocities will be employed than hereinabove indicated and it will likewise be understood that with larger average particle sizes the upward vapor velocities in the system will be necessarily increased. The design of reactors, regenerator chamber, standppes and strippers is of course a matter of engineering skill and such designs can be readily determined for any desired cata'- lyst size. Y

The isoforming reaction may be effected a pressures ranging from atmospheric to 200 i pounds or more per square inch but in our prewas effected at about pounds per square inch l and catalyst was separated from the regeneration gases at about 1 pound per square inch. Two or three pounds per square inch must be allowed for the passage of the catalyst across each of the control valves Ill and |29. It will be readily understood that the whole system may be operated at higher pressures since the standpipes provide the necessary pressure differential to effect the flow of the aerated catalyst through the closed catalyst system.

Steam may be employed in amounts of from 1 to 15 percent or more by weight. 'I'his steam is preferably introduced through line 91 for injecting the powdered catalyst into the transfer line but an additional steam may if desired be introduced into line 98 into the charge entering the pipe still. The steam not only performs the important function of removing the catalyst into the oil stream but promotes high catalyst activity. Steam may also be employed in the regeneration system for removing the heat of regeneration and thus maintaining the desired regeneration temperature.

Our invention has been described by examples illustrating the use of very small catalyst-to-oil ratios and relatively large catalyst-to-oil ratios but it should be understood that any intermediate catalyst-to-oil ratio may be employed provided only that the reactor be designed to insure a contact with the oil with an amount of catalyst which is roughly equivalent to the amount of catalyst that would be employed in the xed bed system described in our application Serial No. 350,270. The'low catalyst-to-oil ratio offers the advantage of longer once-through catalyst holding times in the reactor. With relatively short once-through holding times in the reactor the catalyst will not be fully spent and it is, therefore, desirable to recycle this catalyst prior to regeneration. The larger the catalyst-to-oil ratio and the smaller the catalyst holding time per pass in the reactor, the greater will be the amount of catalyst recycled as compared to the amount regenerated.

In both preferred examples we have illustrated systems wherein the treating ratio is about 100 barrels of feed stock per hour per ton of catalyst present in the reactor. The solid line curves in Figures 3 and 4 are based on coker naphtha, with coil and drum cracked naphtha the curve in Figure 4 will have approximately the shape indicated by the dotted line curve, showing that substantially optimum conversion may be obtained with 1 ton or even 1/2 ton of catalyst per 100 barrels of naphtha feed per hour. For active catalysts the treating rate is preferably in the range of .25 to 2.5 tons of catalyst in the reactor per 100 barrels of naphtha charge per hour for a catalyst residence time of about 24 hours.

The overall catalyst residence time in the reactor is the average length of time that a particle of catalyst is in the reactor between regeneration steps. In dense phase operations with no catalyst recycling, the once through holding or residence time is also the overall residence time. In catalyst recycling operations the overall holding or residence time is the sum of the once-through holding times of an average catalyst particle between regeneration steps. Thus rii the once-through holding time is 10 seconds and only .07% of the catalyst is by-passed for regeneration, an average catalyst particle will make about 1430 passes of 10 seconds each through the reaction chamber between regeneration steps, giving an overall residence time of about 4 hours. With short overall catalyst residence time in the reactor, however, lower total amounts of catalyst need be present in the reactor because of 4the greater catalyst activity and the lesser amounts of carbonaceous deposit thereon. It may thus be possible to operate a powdered catalyst system with .05 to .005, for example .02 ton of catalyst in the reactor per 100 barrels of thermally cracked naphtha charged per hour by virtue of very small overall holding times. The time of oil contact may vary over a wide range, for example from about 2 to 20 seconds or more. The vapor velocities, reactor designs, etc. may be calculated by those skilled in the art in order to eiect the isoforming under the conditions hereinabove set forth. The relationship of overall catalyst holding time to the amount of catalyst used may be expressed by the following equation:

T=at-m where: T is tons of catalyst in the reactor per 100 barrels of thermally cracked naphtha feed per hour, a is a constant in the range .05 to .005, for example .02, and t is the overall catalyst residence time in minutes.

thereof and that the invention is applicable to a fairly wide range of operating conditions and procedural steps. The isoforming step of our invention is applicable only, however, to thermally cracked naphthas, preferably having an end point not higher than about 400 to 450 F. and it is -not applicable to catalytically cracked naphtha nor to virgin naphtha. cracking step is preferably at high temperature and low pressure because thermally cracked naphtha so formed gives a much greater octane number improvement on isoforming than does a naphtha produced by thermal cracking at high pressure and high temperature or at any pressure and low temperatures.

We claim:

l. The method of producing high octane number motor fuel from charging' stock of the class consisting essentially of gas oil and heavier hydrocarbons which method comprises thermally The thermal cracking said charging stock to produce a thermally cracked naphtha having an end point below about 450 F. together with lighter and heavier products, removing the heavier products from the thermally cracked naphtha, heating the thermally cracked naphtha to a temperature of about 850 to 1100 F., introducing into a contacting zone a powdered isoforming catalyst and the hot naphtha stream in amounts suiiicient to give a catalyst-to-oil weight ratio of materials entering said contacting zone within the approximate range of .001:1 to 10:1, retaining catalyst in said contacting zone in amounts and for a period of time as defined by the following formula:

where T is tons of catalyst in contacting zone per hundred barrels of charging stock per hour which is introduced thereto, a is a constant ranging from about .05 to about .005 and t is the overall catalyst residence time in the reactor in minutes, separating the powdered catalyst from reaction vapors leaving said contacting zone and fractionating the vapors to obtain a gasoline fraction, a heavier-than-gasoline fraction and a lighter-than-gasoline fraction.

2. The process of claim 1 wherein the oil contact time in the reaction chamber is about 2 to 20 seconds.

3. 'I'he method of claim 1 'wherein the pressure in the contacting step is about atmospheric to 50 pounds per square inch and wherein the temperature is about 900 to 975 F.

4. The method of claim 1 wherein the catalystto-oil Weight ratio of materials entering said ,I

contacting step is within the approximate range of .006:1 to .02:1 and wherein the major part' of the catalyst is regenerated between contacting steps.

5. The method of claim 1 wherein the catalystto-oil weight ratio of materials entering said contacting step is within the approximate range of about .02:1 to 10:1 and wherein the major part of the catalyst is recycled to said contacting step without regeneration.

6. The method of producing high octane number motor fuel from charging stock'of the class consisting essentially of gas oil and heavier hydrocarbons which method comprises thermally cracking said charging stock to produce a thermally cracked naphtha having an end point below about 450 F. together with lighter and heavier products, removing theheavier products from the thermally cracked naphtha, heating the thermally cracked naphtha to a temperature of about 850 to 1100 F., injecting a powdered isoforming catalyst into a stream of hot naphtha vapors at said temperature, conducting said catalyst by means of said stream to an enlarged reaction zone containing 2.5 to 0.1 tons of catalyst per barrels of thermally cracked naphtha charged per hour, regulating the vapor velocity in said reaction zone to maintain a layer of dense phase catalyst therein and to obtain a time'of contact of about 2 to 20 seconds between naphtha vapors and said dense phase catalyst, separating catalyst from vapors leaving said reaction chamber, stripping said catalyst with inert gas, pneumatically conveying said spent catalyst to a regeneration zone, introducing air at the base of said regeneration zone in amounts sufficient to burn carbonaceous material from the catalyst, maintaining a dense catalyst phase in said regeneration 'zone by controlling vapor velocity therethrough,

separating catalyst from regeneration gases leaving said regeneration zone, stripping regeneration gases from regenerated catalyst and reintroducing regenerated catalyst into the hot naphtha vapor stream.

'1. The method of increasing the octane number of thermally cracked naphtha which comprises vaporizing and heating said naphtha to a temperature of about v850 to 1025 F., treating said vaporized naphtha in a contacting zone with an amount of isoforming powdered catalyst dened by the following formula where T is tons of catalyst in the contacting zone per hundred barrels of charging stock per hour introduced into said contacting zone,a. is a constant ranging from about .05 to .005 and t is overall catalyst residence time in minutes, maintaining a catalyst-to-oil weight ratio of materials introduced into the contacting zone within the ap. proximate range of .001:1 to 10:1, regenerating spent catalyst, returning regenerated catalyst to said contacting zone, separating catalyst from r vapors leaving said contacting zone, and fractionating said vapors to obtain therefrom a fraction boiling in the gasoline boiling rrange.

8. The method of claim 7 wherein the catalystto-oil ratio entering said contacting zone is within'the approximate range of .001:1 to .02:1 and wherein the major part of the catalyst is regenerated between treating steps.

9. The method of claim 7 wherein the catalystto-oil weight ratio of Imaterial entering said contacting zone is within the approximate range of about .02: 1 to about 10:1 and wherein the majorl part of the catalyst separated from the contacting zone is recycled to said contacting zone without regeneration.

10. The method of improving the octane number of a thermally cracked naphtha which comprises contact-ing said naphtha with a siliceous metal oxide catalyst suspended in a vapor stream at a temperature of about 800 to 1100 F. and with an amount of catalyst dened by the following equation:

T=at0.534

Where T is tons of catalyst in the reaction zone per 100 barrels of charging stock per hour, a is a constant ranging from about .05 to .005, and tis the overall catalyst residence time in minutes, and Where the catalyst-to-oil weight ratio of materials entering the reaction zone is within the approximate range of .001: 1 to 10: 1.

11. The method of producing high octane number motor fuel from thermally cracked naphtha which method comprises vaporzing and heating said naphtha to a temperature of about 850 to 1100" F. and introducing said vaporized naphtha at the base of a substantially vertical contacting zone containing 2.5 to 0.1 tons of powdered siliceous, metal-oxide catalyst per hundred barrels of naphtha charged per hour, the particle size of the catalyst being in the general vicinity of 20 to 80 microns, maintaining an upward vertical vapor flow through said contacting zone within the general vicinity of about one and one-half feet per second whereby the catalyst is mainaction zone, contacting said naphtha vapors with said catalyst in said reaction zone for a period of time within the general vicinity of 2 to 20 seconds, maintaining substantially constant cattained in dense phase in the lower part of the re- 15 alyst activity in the contacting zone by employing a regenerated catalyst-to-oil weight ratio of materials entering the reaction zone ranging from about 0.1:1 to about .001:1, removing catalyst and vapors from the reaction zone at substantially the same rate as catalyst and vapors are introduced thereto and separating catalyst from reaction products.

ERNEST W. THIELE.

GEORGE E. SCHMITKONS.

CARL MAX HULL. 

